Natural gas resources are located in many areas which are remote from means for transporting such natural gas conveniently and/or economically to a market. In many remote locations the natural gas is co-produced with crude oil and must be disposed of, by flaring or reinjection, in order to produce the crude. Flaring has become an unacceptable disposal method since it wastes a diminishing hydrocarbon resource and is also a source of air pollution. Reinjection, which adds to the cost of crude oil production, is often unacceptable both in view of its cost and the adverse effects it may impose upon crude oil production from the field itself. The inability to dispose of natural gas produced in association with crude at a remote location in a manner which is economically, governmentally and environmentally acceptable has brought crude oil production at some locations to a halt.
Carbon containing compositions, such as coal and natural gas, may be converted to other useful hydrocarbon forms by first converting the carbon composition to a synthesis gas. A synthesis gas is one containing at least hydrogen (H.sub.2) and carbon monoxide (CO). A synthesis gas may be reacted over a variety of catalysts under various conditions to cause the H.sub.2 and CO content thereof to react to form a variety of carbon containing compounds ranging from methanol (CH.sub.3 OH), dimethyl ether (DME), normally liquid hydrocarbons, etc. Methods for the production of synthetic gasoline and diesel fuel, whether by Fischer-Tropsch technology or by the Methanol to Gasoline (MTG) technology developed by Mobil, all require the production of a synthesis gas.
Two basic methods are employed to convert a carbon source to a synthesis gas, namely steam reforming or by adiabatic reaction of the carbon with an amount of oxygen less than the stoichiometric quantity required for complete carbon oxidation. Adiabatic reforming is the only possible method for materials boiling higher than naphtha. For natural gas and heavier material, up to naphtha, steam reforming (commonly over a nickel-containing catalyst) is the preferred method. Especially for ammonia production, but also in other cases, adiabatic reforming over such a nickel catalyst is also often practiced. The preparation of a synthesis gas by steam reforming non-adiabatically, that is with a large heat input, is a process attendant with large capital and high operating cost. Partial oxidation with or without catalyst present produces a synthesis gas of a lower hydrogen content than does steam reforming. In order to avoid the introduction of large amounts of nitrogen (N.sub.2) as an inert diluent into the system (which is a desire for all targeted uses of synthesis gas, with the exception of production of ammonia), essentially pure oxygen must be used. Accordingly, production of an essentially nitrogen free synthesis gas by adiabatic reforming is also an intensive capital and operating cost process since an oxygen separation unit is required together with compressors of special construction required for the safe compression of gases containing 35 mole % or greater oxygen. Adiabatic reforming is generally employed in order to provide an operator flexibility for using different carbon feedstocks ranging form natural gas to readily volatizable hydrocarbon compounds.
Many proposals have been set forth for processes which would recover remote location natural gas in a normally liquid hydrocarbon form by converting it on location to a synthesis gas and then processing such synthesis gas by known methods to convert it to methanol or to other liquid hydrocarbon forms. Since the world market for methanol is insufficient to accommodate that quantity of methanol which would result by recovery of remote natural gas in this form, this manner of resolving the problem has not yet been practiced.
With the development in the mid 1970's by Mobil of a catalytic process for the conversion of methoxy containing compounds, such as methanol and DME, to gasoline grade liquid hydrocarbons, it appeared conceivable to recover remote natural gas on site in the form of a normally liquid hydrocarbon.
Such processes are intensive in their capital and operating cost, in large part due to the manner by which the natural gas is reformed to a synthesis gas and the need to match the synthesis gas pressure to the conditions required for its conversion to methanol and then to gasoline in the Mobil MGT process.
Nevertheless, as crude oil prices rose dramatically during the 1970's and sustained its high level in the 1980's it appeared that gasoline produced from natural gas which would economically compete with refined gasoline could be accomplished by coupling a conventional methanol production plant front end to a Mobil MTG process as the finish end.
In the early 1980's New Zealand, which then depended for its gasoline supply totally on imported crude oil products, undertook at a cost of about 1.2 billion dollars to construct a plant for production of gasoline from methane. The overall plant design comprised two main units, one for the production of methanol from methane, and the second using the Mobil MTG technology for converting methanol to gasoline. In effect, the New Zealand synthetic gasoline plant is two separate plants built side-by-side on common grounds.
Installation of the New Zealand plant was complete and operations commenced in 1985. At that time, crude oil prices had fallen significantly from their previous level and synthetic gasoline produced by the New Zealand plant was, and still is, economically uncompetitive with the price of refined gasoline; in major reason because of the cost, both capital and operating, associated with producing methanol from methane.
In an attempt to improve the economies of synthetic gasoline production using the Mobil MTG process Haldor Topsoe developed a process now commonly known as the Tigas process. The Tigas process integrates methanol synthesis and gasoline synthesis into a single process loop which eliminates the separation of methanol as a discrete intermediate product. To accomplish this integration, Tigas combines both strains of conventional wisdom prevailing in standard methanol production operations in order to eliminate the need to compress synthesis gas from a steam reformer to the pressure required for operation of a methanol plant. Accordingly, in the Tigas process, methane is first steam reformed in part at a pressure of about 30 to 50 atmospheres (440-730 psi) to a high CO.sub.2 content precursor synthesis gas and the unreacted methane content of this precursor synthesis gas is then secondarily reformed by partial oxidation with essentially pure oxygen to produce a still CO.sub.2 rich final synthesis gas having a pressure of about 28 to 48 atmospheres (410-700 psi). This final moderate pressure synthesis gas is then sent to a reactor containing a catalyst which is active for producing both methanol and dimethyl ether from the synthesis gas. Although this reactor operates at a somewhat lower pressure than does a methanol only reactor, because of its coproduction of dimethyl ether a high conversion of methane based carbon to combined methanol and dimethyl ether is still obtained. Total conversion of natural gas input carbon to a methoxy compound containing feed stream composition upon which the Mobil MTG process can operate is high. The methanol and dimethyl ether containing product gas stream is then reacted over a Mobil catalyst to convert the methoxy compounds thereof to liquid hydrocarbon compounds which are separated from the product gas stream and a portion of the residual overhead gasses containing unreacted hydrogen, carbon dioxide, methanol, ethane and olefins are recycled back to the inlet of the methanol/dimethyl ether reactor.
Although the Tigas design somewhat improves the economics for synthetic gasoline production from methane, it still requires a high capital cost steam reforming unit to which Tigas adds a requirement for a high capital cost oxygen plant to permit secondary reforming. The high capital cost required for a synthesis gas compressor is eliminated by Tigas in favor of a high capital cost oxygen plant to obtain in the tradeoff, a net reduction of capital and operating cost, after the obtainment of the synthesis gas, in the form of units of smaller duty size down stream. Though an improvement, given the current price for crude oil, the Tigas process is not economically feasible for synthetic gasoline production from methane in light of its high attendant capital cost.
Some variations to the basic Tigas process have been reported to further reduce the need for high capital cost items. One such variation is reported in U.S. Pat. No. 4,481,305. In this variation, an improvement in the economies of recycle is reported to be obtained compared to the standard recycle procedure described by U.S. Pat. No. 3,894,102 to be used with the Mobil MTG process. The improvement requires that adjustments be made to the composition of the synthesis gas feed to a methanol/dimethyl ether production reactor such that the synthesis gas feed will contain carbon monoxide and hydrogen in a CO/H.sub.2 ratio of above 1 and contain carbon monoxide and carbon dioxide in a CO/CO.sub.2 ratio of from 5 to 20. A synthesis gas of such composition may be obtainable from coal or a similar carbonaceous starting material. It is, however, not economically feasible to prepare a synthesis gas of such composition from methane using the Tigas process.
Even though synthetic gasoline production processes such as Fischer-Tropsch synthesis (FTS), standard Mobil and/or Tigas have undergone steady improvements intended to render them more economical to the production of synthetic gasoline from methane, they are today still unable to produce gasoline at a cost competitive to that refined from petroleum crude. This is so even where a source of low cost methane is conveniently located to or transportable to the synthetic gasoline production plant site.
Application of a currently existing process for conversion of remotely located natural gas to methanol and/or for synthetic gasoline production from natural gas is not economically feasible in view of the great capital cost associated with the equipment necessary to practice such processes.
In commonly owned copending U.S. patent application Ser. No. 508,928, a process is disclosed for converting natural gas to a synthesis gas using a low grade oxygen source, i.e., containing 50% or more nitrogen, which significantly lowers the capital cost of a methane to methanol, dimethyl ether or gasoline production plant. In one embodiment of the disclosed invention a gas turbine is utilized to power the compressors needed to compress methane and air in the process steps of the methane to synthesis gas conversion. In addition, the gas turbine provides a ready supply of compressed air in the range of about 8 to 16 atmospheres absolute. Compressed air from the turbine is bled off in an amount not exceeding the mass balance requirements of the turbine. The so compressed gas is then compressed in a secondary compressor, which may be powered by the energy output of the gas turbine, to the required 400 to 2000 psig oxygen gas compression requirement of the process. Thus, a primary compression of the oxygen containing gas of from about 8 to about 16 fold produced by the turbine is achieved at little or no cost. The remaining 1.8 to 17 fold compression required of the secondary compressor can be achieved at considerable savings over the 28 to 137 fold compression which in the absence of the gas turbine would be performed on the oxygen containing gas by conventional gas compressor units.
As further described in copending U.S. Ser. No. 508,928, the economics and efficiencies of the process of converting natural gas to a recoverable normally liquid hydrocarbon can be further improved by preparing the synthesis gas with an oxygen enriched gas having up to 50 mole % oxygen.
Such oxygen-rich gas normally contains nitrogen as its other main component. Presently there are three different ways to arrive at such a gas, normally described as oxygen-enriched air.
Smaller amounts of such gas are made by preferential diffusion of oxygen over that of nitrogen through an appropriate membrane. In order to drive that diffusion, air is compressed and then fed to the diffusor unit. This diffusor unit can employ flat membranes, but it is often preferred to use a large number of small hollow membrane fibers. Either the compressed air is fed to the inside or to the outside of the membrane fibers. When fed to the inside, the fiber can be thin-walled. This helps the diffusion, but limits the maximum pressure to about 150 psig. It is also possible to feed the compressed air to the outside of the fibers. Then the wall of these fibers has to be thick enough to withstand the pressure. This leads to the use of higher pressures, but the greater wall thickness slows the rate of diffusion down. It should be clear, that the air compression is a cost factor, but higher compression within the limits of tolerance speeds up the diffusion, thus lowering the cost of the fiber material. In practice optimization of these two effects has led to compression of air to 8 atmosphere gage (atg), followed by a rather intensive use of the amount of air. The residual gas, or "spent air" commonly contains little more than 7% oxygen. The cost of the air compression, both in capital and in operation, is substantial. The cost of the membrane fiber material, together with a containment vessel therefor and the necessary air filters, is mostly capital only. Enriched air made this way has a very high oxygen equivalent cost, well over $50 per metric ton. The term "oxygen equivalent cost" derives from the assumption that the same amount of enriched air can also be made by adding pure oxygen to air. When assuming air to have no cost, the cost of the enriched air is the cost of that amount of pure oxygen which must be added to produce an equivalent volume.
While a membrane diffusion method is commonly used for preparation of small amounts of such enriched air, larger quantities may be produced by pressure swing absorption (PSA). Here compressed air is contacted with an absorbent that preferentially absorbs oxygen. As soon as the absorbent is saturated to a reasonable degree with oxygen, the flow of the pressurized air is stopped and the pressure reduced to atmospheric and oxygen is desorbed. Thus mainly oxygen is produced, be it at low pressure. The term pressure swing derives from the alternating of pressurized absorption and low pressure desorption.
The cost of oxygen made by PSA, is much lower than with diffusion. However, the PSA cost is still substantial. For still larger amounts of oxygen, say 500 metric tons per day (MTPD) and higher, cryogenic separation of oxygen is the preferred method. Then the cost of oxygen at present can be brought down to the zone of $25 to a $35 per metric ton.
Hegarty U.S. Pat. No. 4,545,787 describes a process wherein an oxygen enriched gas may be prepared at low utility cost--for the volume produced--by the incorporation of an oxygen permeable membrane separation unit or PSA with the operation of a gas turbine operated for power generation. Unfortunately, in view of the mass and thermal balance tolerance constraints which must be observed for proper operation of the gas turbine, the process described by Hegarty only provides for the separation of a minor amount of oxygen from a portion of the compressed air provided by the turbine.
Feeding air at 11 atmospheres absolute (ata) to a membrane with a gas separation factor for O.sub.2 versus N.sub.2 of 5.55 can provide a permeate stream of about 54% O.sub.2 in a volume amount of about 1.5% of the volume of feed air. The same membrane can produce a 49 mole % O.sub.2 permeate in a volume amount of about 17.6% of the volume of the feed air, or a 40 mole % O.sub.2 permeate of about 40.5% of the volume of the feed air. In the last case the remaining O.sub.2 in the used air amounts to only 8% . In each case the balance of the O.sub.2 enriched permeate gas is comprised essentially of N.sub.2. Theoretically, if all O.sub.2 of the feed air is extracted into the permeate the permeate will be about 35 mole % O.sub.2 and be of a volume of about 50% that of the feed air.
Hegarty U.S. Pat. No. 4,545,787 proposes to compensate for the disruption to the thermal balance between the compressor and expander side of a gas turbine which is caused by the recovery of O.sub.2 from the compressed air by increasing the amount of air compressed by the compressor side in a quantity equal to the amount of O.sub.2 diverted into the permeate O.sub.2 enriched gas stream. Hegarty proposed that this maintains the same molar flow to the expander side as would be the case had no O.sub.2 been recovered.
Turbines are designed to pass all air compressed in the compressor side to the expander side. In normal operation the only mass imbalance between compressor to expander side is due to the mass increase caused in the expander side by reason of fuel supplied to the combustion chamber of the expander side. The turbine design provides for this positive mass imbalance and further for a safety factor of a positive mass imbalance preferably no greater than a 10% increase of mass in the expander side. The operation of a turbine in a manner which exceeds the positive mass imbalance safety limit, or which produces a negative imbalance by decreasing mass flow in the expander side, will result in a significant lessening of its service life and even to its total failure.
Even were one able to put Hegarty's proposal to practice, for the same amount of fuel input to cause the compressor side of the turbine to compress additional quantities of air equal in an O.sub.2 amount to the O.sub.2 recovered, Hegarty's method would not maintain the turbine in a tolerable mass balance even if the requirement of thermal balance were met. Hegarty simply does not address nor consider the problem of proper mass balance.
For example, with a membrane having a gas separation factor for O.sub.2 of 5.55 employed in Hegarty's method with a turbine compressing 1000 lb.-mole/hr air (21% O.sub.2, 79% N.sub.2) which is 314% of that amount required for fuel combustion (33.44 lb.-mole/hr CH.sub.4), without Hegarty's proposed additional air compression the mass imbalance of the turbine would be negative and would exceed 10% when the O.sub.2 recovery reaches or exceeds about 64.3 lb-mole/hr O.sub.2 since the permeate would also contain about 48.7 lb-mole/hr N.sub.2. Following Hegarty's proposal to compress an additional quantity of air equal to the extracted O.sub.2 would produce a negative mass imbalance of 16%; compression of an additional quantity of air equal to the sum of the O.sub.2 and N.sub.2 lost in the permeate would produce a negative mass imbalance of 19.5%.
It is evident that one of two realities will quickly become apparent to one who attempts to practice the process of Hegarty, namely that either the quantity of O.sub.2 produced must be limited to a small quantity to prevent an intolerable mass imbalance or the turbine must be sacrificed to a short service life or even failure.
Accordingly, even in view of Hegarty's proposal, the fact remains the same today that membrane O.sub.2 enrichment is economically viable only for production of small quantities of O.sub.2 ; for medium volume amounts of produced O.sub.2 a pressure-swing absorption method of O.sub.2 production is viable and wherein large quantities of O.sub.2 are required an efficient but capitally intensive cryogenic method. for O.sub.2 production is still the most economically viable process for production of such large quantities.
There is still a need for a method of production of useful O.sub.2 in large volumes at a production cost significantly less than that available by conventional methods. Such a method would allow the use of an oxygen enriched gas for the improvement of many industrial scale processes to provide valuable products at significantly reduced production cost.